Natural gas liquids upgrading process: two-step catalytic process for alkane dehydrogenation and oligomerization

ABSTRACT

A process to catalytically transform natural gas liquid (NGL) into higher molecular weight hydrocarbons includes providing an NGL stream, catalytically dehydrogenating at least a portion of the NGL stream components to their corresponding alkene derivatives, catalytically oligomerizing at least a portion of the alkenes to higher molecular weight hydrocarbons and recovering the higher molecular weight hydrocarbons. The NGL stream can be extracted from a gas stream such as a gas stream coming from shale formations. The higher molecular weight hydrocarbons can be hydrocarbons that are liquid at ambient temperature and ambient pressure.

CROSS-REFERENCE TO RELATED APPLICATIONS

The current application claims priority to U.S. Provisional Patent No.62/555,601 filed on Sep. 7, 2017.

BACKGROUND

The present disclosure is directed to new hydrocarbon upgradingprocesses that include the dehydrogenation of light alkanes to theirrespective alkene derivatives, commonly known as olefins, followed byoligomerization of the light olefins to higher molecular weighthydrocarbon derivatives. The longer chain higher molecular weighthydrocarbons have a greater value than the lighter alkanes.

More efficient utilization of petroleum and gas reserves is an importantstrategy for the deployment of future energy generation. Shale gas hasbecome an increasingly important source of natural gas in the UnitedStates, and the U.S. government's Energy Information Administrationpredicts that by 2040, seventy percent of the United States' natural gassupply will come from shale gas. Many of these shale gas formationscontain wet gases, which can include substantial concentrations ofnatural gas liquids (NGL). NGL is a mixture of hydrocarbons made upprimarily of ethane (C₂), propane (C₃), butane (C₄), and pentane (C₅).Several major shale gas formations such as Marcellus and Bakken arelocated far away from historically gas producing and processing regionsuch as the Gulf Coast. These resources can be considered as strandedgas. Also, associated gas, which is a byproduct of shale oil productionand may contain substantial concentration of NGL, is generally flared,vented, or injected back to the shale oil reservoir at high cost.Construction of pipelines to transport natural gas liquids to largeexisting processing plant complexes, such as the Gulf Coast, can becapital intensive. This creates an opportunity to upgrade this strandednatural gas, particularly its condensate, or NGL, into liquid fuel foruse or further processing as it is easier to transport and distribute tothe market.

The transformation of shale gas to higher molecular weight products suchas transportation fuels, fine chemicals and polymers is one of thestrategies to utilize the shale gas reserves to their highest value.

Currently, UOP Cyclar, Synfuels International ETG (Ethane to Gasoline),and Greyrock DFP (Direct Fuel Production) processes are examples ofcommercialized NGL to Liquids (NTL) processes. The Greyrock DFP processutilizes a small scale Fischer-Tropsch (FT) type process. Unlike atraditional FT process, their proprietary catalyst is claimed toeliminate the need for downstream liquid product upgrading. SynfuelsInternational, Inc. employs a thermal cracking reactor followed by aseries of proprietary reactors that ultimately yields gasoline blendstock as the final product. In the UOP Cyclar process, liquefiedpetroleum gas containing mainly propane and butane is converted toaromatics through dehydrogenation and subsequent aromatization.Aromatics, however, may not be suitable as liquid fuel and its marketshare is smaller than that of liquid fuel. Therefore, upgrading naturalgas liquids into liquid fuel, having a minimal aromatic content, canincrease the value of natural gas liquids and enable the distribution ofthe natural gas liquids into the liquid fuel market.

Currently, dehydrogenation of ethane is typically accomplished usingsteam cracking. In this process, ethane is subjected to high temperature(˜800° C.) and mixed with steam. The resulting vapor contains ethylene,ethane, methane, acetylene, hydrogen, and steam. Typically, steamcracking vapor undergoes cryogenic separation, at great cost, to recoverthe olefins and recycle the unconverted alkanes. In addition, steamcracking is usually implemented at large scale; it is unlikely to beeconomical for small-scale operations in remote shale gas formationregions. UOP Oleflex and ThyssenKrupp STAR processes are also employedfor only propane dehydrogenation using their respective proprietarycatalysts. Currently, applicants know of no technology to uniformlydehydrogenate a mixture of one or more of C₂, C₃, C₄, and C₅ without theuse of thermal cracking.

Upgrading stranded NGLs may facilitate its marketability anddistribution into the liquid fuel market for use or into a refinery forfurther processing. Besides UOP Cyclar, Synfuels International ETG(Ethane to Gasoline), and Greyrock DFP (Direct Fuel Production)processes, no other major NGL-to-Liquid process has been commercialized.

BRIEF DESCRIPTION OF THE DRAWINGS

The accompanying drawings are incorporated into and form a part of thespecification to illustrate aspects and examples of the presentdisclosure. These figures together with the description serve to explainthe general principles of the disclosure. The figures are only for thepurpose of illustrating examples of how the various aspects of thedisclosure can be made and used and are not to be construed as limitingthe disclosure to only the illustrated and described examples.

FIG. 1 is a simplified block flow diagram illustrating how etheneoligomerization forms an entry step chemistry to form higher molecularweight compounds.

FIG. 2 is a block flow diagram for an embodiment of the presentdisclosure.

FIG. 3 is a block flow diagram for an embodiment of the presentdisclosure.

FIG. 4 is a block flow diagram for an embodiment of the presentdisclosure.

FIG. 5 is a block flow diagram for an embodiment of the presentdisclosure.

FIG. 6 is a block flow diagram for an embodiment of the presentdisclosure.

FIG. 7 is a block flow diagram for an embodiment of the presentdisclosure.

DETAILED DESCRIPTION

The following detailed description illustrates embodiments of thepresent disclosure. These embodiments are described in sufficient detailto enable a person of ordinary skill in the art to practice theseembodiments without undue experimentation. It should be understood,however, that the embodiments and examples described herein are given byway of illustration only, and not by way of limitation, the embodimentsare chosen and described so that others skilled in the art mayappreciate and understand the principles and practices of the presentteachings. Various substitutions, modifications, additions, andrearrangements may be made that remain potential applications of thedisclosed processes. Therefore, the description that follows is not tobe taken as limiting on the scope of the appended claims. In particular,an element associated with a particular embodiment should not be limitedto association with that particular embodiment but should be assumed tobe capable of association with any embodiment discussed herein.

Definitions

For the purpose of this description and appended claims, the followingterms are defined.

The use of the terms “a” and “an” and “the” and similar referents in thecontext of describing the elements (especially in the context of thefollowing claims) are to be construed to cover both the singular and theplural, unless otherwise indicated herein or clearly contradicted bycontext. Recitation of ranges of values herein are merely intended toserve as a shorthand method of referring individually to each separatevalue falling within the range, unless otherwise indicated herein, andeach separate value is incorporated into the specification as if it wereindividually recited herein. All processes described herein can beperformed in any suitable order unless otherwise indicated herein orotherwise clearly contradicted by context. The use of any and allexamples, or exemplary language (e.g., “such as”) provided herein, isintended merely to better illuminate the embodiments and does not pose alimitation on the scope of the claims unless otherwise stated. Nolanguage in the specification should be construed as indicating anynon-claimed element as essential.

The term “alkane” or “paraffin” means substantially saturated compoundscontaining hydrogen and carbon only, e.g., those containing <1% (molarbasis) of unsaturated carbon atoms. The term alkane encompasses C₁ to C₆linear, iso, and cyclo alkanes.

As used herein, an “alkene” or “olefin” refers to any unsaturatedhydrocarbon containing one or more pairs of carbon atoms linked by adouble bond. The olefins described herein include cyclic or aliphaticolefins, and include mono-olefins, di-olefins, tri-olefins, etc.

As used herein, a “bimetallic catalyst” is a catalyst having at leasttwo metal components. The term does not limit the number of metalcomponents to only two. The two metals are, at least partially presentin the metallic phase and/or in a metallic alloy state.

The term “C_(n)” hydrocarbon wherein n is a positive integer, e.g., 1,2, 3, 4, or 5, means hydrocarbon having n carbon atom(s) per molecule.The term “C_(n+)” hydrocarbon wherein n is a positive integer, e.g., 1,2, 3, 4, or 5, means hydrocarbon having at least n carbon atom(s) permolecule. The term “C_(n−)” hydrocarbon wherein n is a positive integer,e.g., 1, 2, 3, 4, or 5, means hydrocarbon having no more than n numberof carbon atom(s) per molecule.

As used herein, a “catalyst” is any substance or material which changesthe rate of conversion of alkanes to alkenes but is not, itself,consumed.

The terms “comprise,” “comprises,” “comprising,” “include,” “includes,”“including,” “have,” “haves,” and “having” are interchangeable and notintended to be limiting.

The term “hydrocarbon” means compounds containing hydrogen bound tocarbon, and encompasses (i) saturated hydrocarbon, (ii) unsaturatedhydrocarbon, and (iii) mixtures of hydrocarbons, including mixtures ofhydrocarbons (saturated and/or unsaturated) having different values ofn.

The term “liquid hydrocarbon” means hydrocarbons that are liquid at roomtemperature and ambient pressure.

The term “support” or “substrate” as used herein is not meant toindicate that this component is necessarily inactive, while the othermetals and/or promoters are the active species. On the contrary, thesupport or substrate can be an active part of the catalyst. The termsubstrate would merely imply that the substrate makes up a significantquantity, generally 10% or more by weight, of the entire catalyst.

The term “zeolite” means microporous, crystalline silicon oxide mineralscommonly used as commercial adsorbents and catalysts.

The present disclosure relates generally to new hydrocarbon upgradingprocesses that include the dehydrogenation of light alkanes to theirrespective alkene derivatives, commonly known as olefins, followed byoligomerization of the light olefins to higher molecular weighthydrocarbon derivatives. The longer chain higher molecular weighthydrocarbons have a greater value than the lighter alkanes. Variousembodiments to the two-step process are disclosed along with additionalelements that can be implemented such as efficiency improvements andintegration into existing processes and/or facilities.

Dehydrogenation of NGLs followed by oligomerization of the resultingalkenes may have the potential as an economically attractive process.This process is termed as Natural Gas Liquid-to-Liquid (NTL). In the NTLprocess, a mixture of light alkanes such as ethane, propane, butane, andpentane and combinations thereof are activated by catalyticallytransforming them into their corresponding olefins. Then, withoutseparation of the olefins from the unconverted alkanes, the olefins arethen catalytically oligomerized to form higher molecular weight liquidhydrocarbons. The resulting hydrocarbons may still contain olefins andcan optionally have a final saturation step to fully saturate thehydrocarbon liquids. The resulting hydrocarbons can then be recoveredand can serve as drop-in liquid fuels, as a refinery feedstock, asfeedstock for chemical plants or other uses.

One aspect of the invention is that, instead of steam cracking,catalytic dehydrogenation is adopted in the process disclosed herein. Ithas the potential to efficiently transform light alkanes to alkenes. Asstated earlier, steam cracking is commonly implemented at large scaleand it is unlikely to be economical at small scale. In addition, steamin the product stream is detrimental to the oligomerization catalystlife; therefore, requires separation of the steam prior to processing ofthe olefins. Production of olefins produced by alkane dehydrogenationallows for direct oligomerization of the mixed olefin, alkane andhydrogen stream without costly separation. Considering the remotelocations and discounted prices of stranded gas, modular scale processmay be preferable to exploit the economic opportunities presented bystranded natural gas and its condensate and NGL content.

Considering the heat duty alone, steam cracking consumes approximately30% more energy for every ethylene molecule produced compared tocatalytic dehydrogenation and this is mainly due to the heat demand forsteam generation in steam cracking. In addition, given appropriatecatalysts, catalytic dehydrogenation can be more selective towardsalkenes and minimizes the production of by-products, which would callfor the need of either downstream upgrading or removal.

Another aspect of the invention is deploying the described NTL processesusing catalytic dehydrogenation and oligomerization at modular scales.Depending on the location and behavior of the shale gas wells, theeconomic opportunity presented by shale gas NGL's can be exploited ateither large or small scales or both. Modular NTL plants enable localprocessing of NGL from stranded or associated gas, which then can bedirectly sold into the local market and mitigates the expense associatedwith distribution of the liquid products from these modular plants.Process modularization of the described NTL process may be achievedthrough process intensification, process integration, and novel unitoperation designs of the catalytic reactors and other unit operationsdescribed in the processes.

Embodiments of the present disclosure include the addition of an NTLplant to an existing NGL extraction plant. In many instances NGLextraction plants have no market for the ethane, or the economics aresuch that the extraction, transportation, fractionation and marketing ofthe ethane costs more than the ethane can bring, thus a net loss for theNGL extraction plant operator. In these instances the plant can beoperated in an “ethane rejection” mode where a large portion of theethane is sent out with the residue gas and not recovered in the NGLstream. This contained ethane remains in the natural gas stream and issold at methane prices and used as natural gas. This is a tremendouswaste of a natural resource. The addition of an NTL plant to an existingNGL extraction plant can enable the recovery of ethane by operating theNGL plant in an “ethane recovery” mode, having the mixed NGL stream,including the additional ethane recovered, as a feed stream to the NTLplant, and the NTL plant producing a heavier molecular weighthydrocarbon stream that contains little or no ethane content. Optionallythe ethane, with or without higher hydrocarbons, can be separated fromthe NGL stream and fed to an NTL plant, producing a heavier molecularweight hydrocarbon stream that contains little or no ethane content. Theheavier molecular weight hydrocarbon stream can then be blended with theNGL stream to produce a mixed hydrocarbon stream that contains little orno ethane content. In this way the ethane contained in the natural gasstream can be recovered and upgraded by the NTL plant to an economicallyprofitable product. The heavier molecular weight hydrocarbon stream fromthe NTL plant can be further processed in the same manner as the NGLstream prior to the installation of the NTL plant.

Dehydrogenation

The present disclosure includes one or more dehydrogenation reactions ofalkanes to olefins. The catalytic dehydrogenation of hydrocarbonsaccording to the formula C_(n)H_(2n+2)↔C_(n)H_(2n)+H₂, is a highlyendothermic equilibrium reaction, the reaction rate of which is limitedthermodynamically and which depends on the respective partial pressuresand temperature. The dehydrogenation reaction is favored by low partialpressures of the hydrocarbons and by high temperatures.

The dehydrogenation reactions of alkanes to olefins can be performedover conventional catalysts commonly known in the art. Commerciallyavailable light alkane dehydrogenation catalysts currently in use arePt—Sn and CrO_(x) based catalysts supported on an alkali (Na/K) modifiedalumina. A Pt—Sn catalyst is used in the Oleflex process which utilizesa continuous fluidized catalyst bed. The Oleflex process is commerciallyoffered by Honeywell UOP. Alternately a CrO_(x) catalyst is used in theCatofin process licensed by Lummus Technology, a CB&I company, whichuses parallel fixed beds. Both of these catalysts are reported toachieve selectivity above 90%. Nevertheless, frequent regeneration isneeded in industrial operation and the catalysts have a total life ofapproximately 1 to 3 years.

Due to equilibrium limitations, these selective alkane dehydrogenationreactions are typically carried out at a high temperature to maximizethe alkene yield and minimize yields of undesired products such asalkynes, diolefins, and cracking products (lower molecular weight alkaneolefin mixtures). Useful catalysts must exhibit high activity andselectivity for the desired dehydrogenation process and a minimal rateof deactivation. Dehydrogenation reactions are known to produce cokewhich is highly refractory and the coke formation leads to catalystdeactivation. Coke removal can require combustion in oxygen containinggas at temperatures greater than 600° C. Desirable catalysts, therefore,must retain high alkane dehydrogenation activity following hightemperature regeneration.

The dehydrogenation reactions can be performed over novel catalystcompositions. In one embodiment the dehydrogenation catalyst isbimetallic and comprises a combination of one or more Group VIII noblemetal, such as platinum or palladium, and a metal selected from thegroup consisting of manganese, vanadium, chromium and titanium, on asupport. The Group VIII noble metal can be present in an amount rangingfrom 0.001 wt % to 30 wt % on an elemental basis of the catalystcomposition. The manganese, vanadium, chromium, titanium, andcombinations thereof, can be present in an amount from 0.001 wt % to 30wt % on an elemental basis of the catalyst composition and are present,at least partially in the metallic phase. The manganese, vanadium,chromium, titanium, and combinations thereof can, at least partially bepresent in an alloy state. Non-limiting examples of support can includesilicon dioxide, aluminum oxide and titanium dioxide. The catalyst canbe an active and selective catalyst for the catalytic dehydrogenation ofalkanes to olefins while retaining high activity and selectivity evenfollowing repeated regeneration by burning coke in oxygen.

The catalysts of the present disclosure can be prepared by impregnatinga support material with manganese, vanadium, chromium, titanium, orcombinations thereof, to form a precursor. The precursor can then bedried and calcined. The amount of manganese, vanadium, chromium,titanium, and/or combinations thereof, to be impregnated can range fromabout 0.001 wt % to about 30 wt %, optionally from about 1.0 wt % toabout 5 wt % calculated on an elemental basis of the final catalystcomposition.

A metal selected from the group consisting of Group VIII noble metal(s)is then impregnated onto the modified support material to providedehydrogenation functions. The Group VIII noble metal can be selectedfrom the group of platinum, palladium, osmium, ruthenium, iridium,rhodium, or combinations thereof. In an embodiment either platinum,palladium, or combinations thereof are employed. The amount of noblemetal loading to be impregnated can range from about 0.001 wt % to about30 wt % calculated on an elemental basis of the final catalystcomposition. It is desirable that the catalyst will contain from about0.1 wt % to about 5 wt % noble metal, most desirable about 0.3 wt % toabout 2 wt % noble metal. As an example, platinum loading on silica canbe accomplished via incipient wetness impregnation techniques using anaqueous solution of platinum tetraammine nitrate Pt(NH₃)₄(NO₃)₂ having apH adjusted with ammonium hydroxide. Following noble metal loading thecatalyst can be dried, calcined and reduced. Platinum loading on aluminacan be accomplished via incipient wetness impregnation techniques usingan aqueous solution of chloroplatinic acid H₂PtCl₆ having a pH adjustedsolution with HCl.

In an embodiment the dehydrogenation catalyst is a bimetallic catalystcomposition that includes a Group VIII noble metal selected from thegroup consisting of nickel, iron, cobalt, and combinations thereof,along with a metal selected from the group consisting of molybdenum,indium, phosphorous, zinc and combinations thereof, and a support. In anembodiment the Group VIII noble metal can be present in an amountranging from 0.01 wt % to 30 wt % on an elemental basis of the catalystcomposition. The molybdenum, indium, phosphorous, zinc, and combinationsthereof, can be present in an amount from 0.001 to 30 wt % on anelemental basis of the catalyst composition. In an embodiment the metalselected from the group consisting of molybdenum, indium, phosphorous,zinc, and combinations thereof are present, at least partially in themetallic phase.

The bimetallic catalyst composition can be produced by providing asupport material, adding to the support material a Group VIII noblemetal selected from the group consisting of nickel, iron, cobalt, andcombinations thereof, and a metal selected from the group consisting ofmolybdenum, indium, phosphorous, zinc, and combinations thereof, to makea catalyst material, calcining the catalyst material; and reducing thecatalyst material to form a dehydrogenation catalyst.

The metals can be added in any suitable manner known in the art, such asnon-limiting examples of supported on a substrate or an inert support,added to a binder, placed on or within a zeolite or other catalystsupport, such as by ion exchange, incipient wetness impregnation, porevolume impregnation, soaking, percolation, wash coat, precipitation, andgel formation.

The various elements that make up the components for the catalyst can bederived from any suitable source, such as in their elemental form, or incompounds or coordination complexes of an organic or inorganic nature,such as carbonates, oxides, hydroxides, nitrates, acetates, andchlorides. The elements and/or compounds can be prepared by any suitablemethod known in the art for the preparation of such materials.

The active metals individually can range from 0.01% to 20% by weight ofthe catalyst, optionally from 0.1% to 10%. If more than one active metalis combined, they together generally can range from 0.01% up to 30% byweight of the catalyst. The manganese, vanadium, chromium, titanium, andcombinations thereof can, at least partially be present in the metallicphase. The manganese, vanadium, chromium, titanium, and combinationsthereof can, at least partially be present in an alloy state.

The supports of the present disclosure can be any suitable support, suchas for non-limiting examples: silicon dioxide, aluminum oxide, titaniumdioxide, zeolites, silica-alumina, cerium dioxide, zirconium dioxide,magnesium oxide, silica pillared clays, metal modified silica, metaloxide modified silica, silica-pillared clays, metal oxide modifiedsilica-pillared clays, silica-pillared micas, metal oxide modifiedsilica-pillared micas, silica-pillared tetrasilicic mica,silica-pillared taeniolite, zeolite, molecular sieve, and combinationsthereof.

The present disclosure is not limited by the method of dehydrogenationcatalyst preparation, and all suitable methods should be considered tofall within the scope herein. Conventional methods includeco-precipitation from an aqueous, an organic, or a combinationsolution-dispersion, impregnation, incipient wetness impregnation, drymixing, wet mixing or the like, alone or in various combinations. Ingeneral, any method can be used which provides compositions of mattercontaining the prescribed components in effective amounts. Otherimpregnation techniques such as by soaking, pore volume impregnation, orpercolation can optionally be used. Alternate methods such as ionexchange, wash coat, precipitation, and gel formation can also be used.

Binder material, extrusion aids or other additives can be added to thedehydrogenation catalyst composition or the final catalyst compositioncan be added to a structured material that provides a support structure.The combination of the dehydrogenation combined with additional elementssuch as a binder, extrusion aid, structured material, or otheradditives, and their respective calcination products, are includedwithin the scope of the invention.

The prepared dehydrogenation catalyst can be ground, pressed, sieved,shaped, extruded and/or otherwise processed into a form suitable forloading into a reactor. The reactor can be any type known in the art,such as a fixed bed, fluidized bed, or swing bed reactor. Optionally aninert material, such as quartz, alpha-alumina, or others can be used tosupport the catalyst bed and to locate the catalyst within the bed.Depending on the catalyst, a pretreatment of the catalyst may, or maynot, be necessary.

In an embodiment, the disclosure includes a process for thedehydrogenation of alkanes to olefins. The process includes the steps ofintroducing an alkane feedstock into a reaction chamber, passing thefeedstock over a dehydrogenation catalyst at reaction conditionseffective to provide a product containing olefin hydrocarbons, andregenerating the catalyst in-situ, when necessary.

The alkane feedstock can be alkanes containing less than 6 carbon atoms.The feedstock can consist primarily of C₂-C₅ alkanes. An embodiment ofthe invention provides for the use of ethane or propane or butane or amixture of these gases as the starting material. Embodiments of theinvention are particularly suitable to produce ethene or propene orbutene or a mixture of these olefins. The alkane feedstock can beobtained from the side product of various hydrocarbon processing plants,for instance, the offgas of an FCC cracker or other refiners processes,refinery fuel gas, or shale gas hydrocarbons. One source of alkanefeedstock is from NGL's that can be extracted by gas processing plants,often a cryogenic process that extract the NGL's from a gas stream, suchas a gas stream produced from a shale formation. One source of alkanefeedstock is liquid petroleum gas (LPG), which consists mainly of thepropane and butane fraction and can be recovered from gas and oil fieldsand petroleum refining operations. Co-feed can contain hydrogen. In anillustrative embodiment the alkane feed can contain primarily ethane. Inan illustrative embodiment the alkane feed can contain primarilypropane. In an illustrative embodiment the alkane feed can containprimarily butane. In an illustrative embodiment the alkane feed cancontain primarily ethane and propane. In an illustrative embodiment thealkane feed can contain primarily propane and butane. In an illustrativeembodiment the alkane feed can contain primarily butane and pentane. Inan illustrative embodiment the alkane feed can contain primarily C₃-C₅alkanes. In an illustrative embodiment the alkane feed can containprimarily C₄-C₅ alkanes.

The reaction chamber used in the dehydrogenation reaction may use one ormore than one dehydrogenation catalysts. They can also house anysuitable catalyst system, such as a fixed catalyst bed, a moving bed ora fluidized bed. Single or multiple catalyst beds can be used, and thereactor can be a swing reactor. The catalysts described herein may beused in any suitable reactor. The process could utilize a series offixed bed reactors, where each reactor could be independentlyregenerated, a moving bed reactor where the catalysts moves through thereactor and is regenerated in a separate section of the plant, or afluidized bed reactor, where the catalyst is circulated through thereactor and regenerated in a separate vessel.

The dehydrogenation reaction can take place at a temperature of from350° C. to 1000° C., optionally from 400° C. to 800° C., optionally from450° C. to 750° C. The pressure can be in the range of from 3 psig to600 psig, optionally from 3 psig to 300 psig, optionally from 3 psig to150 psig. The weight hourly space velocity can be from 0.3 to 50 hr⁻¹,optionally from 0.3 to 10 hr⁻¹, and optionally from 0.3 to 3 hr⁻¹.

The dehydrogenation reaction can be performed adiabatically ornon-adiabatically or approximately isothermally. To achieve reasonablereaction rates, several catalyst beds are normally arranged in seriesand the reaction system is re-heated downstream of each catalyst bed.

If the dehydrogenation is performed in a non-adiabatically operatedcatalyst bed, the catalyst bed can be heated to maintain elevatedtemperature. Because the temperature in the reaction system is keptconstant, the reaction rates may be kept appropriately high. Because ofthe location of the point of thermodynamic equilibrium, however, thedisadvantage is that these high reaction rates can only be achieved atelevated temperatures, because of which the selectivity of olefinformation may be reduced. Hence, consecutive reactions will increasinglytake place, so that undesired products may form, such as CH₄ and coke.

The by-products thus formed, especially finely dispersed coke, candeposit during the reaction on the catalyst, thus causing its state tochange continually. The catalyst can become coated with an undesiredsubstance and is thus less accessible for the reactants. This means thatthe catalyst becomes deactivated. The activity of the catalyst foralkane dehydrogenation and the selectivity for alkene formation may inturn deteriorate. This would result in deterioration of the efficiencyof the process. To counter-act this negative influence on the process,the catalyst will have to be regenerated after a certain reaction periodto recover its activity.

In an embodiment the dehydrogenation catalyst of the present disclosurecan undergo in-situ regeneration. The regeneration can be done at thereaction temperature by burning of carbon with oxygen concentrationsbetween 0.1-20%, optionally from 0.3-10%, and optionally from 0.5-3%.Alternatively, the catalyst can be regenerated with hydrogen at thereaction temperature. In an embodiment the catalyst of the presentdisclosure can undergo ex-situ regeneration.

Depending on its characteristics, the catalyst can be regenerated bybringing it in contact with an oxygen-bearing gas under conditionsdefined for the regeneration of the catalyst. The conditions for suchregeneration may differ from those required for the dehydrogenation. Anoxygen-bearing gas diluted with steam may also be fed through thecatalyst. Because of this procedure, the by-products on the catalyst arereduced, with the result that the catalyst can regain its activity. Ifan oxygen-bearing gas diluted with steam is used for catalystregeneration, the carbon-bearing deposit reacts to form carbon dioxideas the main product. The carbon-bearing deposit is converted to gaseousproducts by this reaction and is removed from the system.

As the conditions for the alkane dehydrogenation process differ from thecatalyst regeneration process, the alkane dehydrogenation process willbe interrupted after a certain period of operation and substituted bythe catalyst regeneration process. Thereafter, the reactor bed is purgedand again made available for dehydrogenation. Both these processes, i.e.the alkane dehydrogenation and catalyst regeneration, are thus performedperiodically. To render the overall process economically efficient, thiscan take place in two or a plurality of catalyst beds, in which thereaction and regeneration processes are alternately implemented. Toensure optimum catalyst regeneration, the regeneration process can beinstrumented, monitored and controlled. For fixed bed reactors, aplurality of reactors in parallel can be employed to enable theconcurrent regeneration of the reactors while other reactors are stillin use.

The dehydrogenation reaction products can then be further processed,such as in an oligomerization reaction.

Oligomerization

The present disclosure includes one or more dimerization and/oroligomerization reactions of light olefins. More particularly, thepresent disclosure relates to catalysts which can enable oligomerizationof light olefins to longer chain olefin derivatives.

The oligomerization of light olefins, such as alkene molecules havingfrom 2 to 5 carbon atoms, is an important industrial reaction andrepresents a route to the production of intermediates used to producemotor fuels, plasticizers, pharmaceuticals, dyes, resins, detergents,lubricants and additives. The oligomerization of light olefins, such asethene and propene, represents an important industrial route to theproduction of environmentally friendly synthetic liquid products,especially diesel fuels substantially free of sulfur and aromatics.Thus, ethene oligomerization forms an entry step chemistry to formhigher molecular weight compounds, which is shown in the block flowdiagram of FIG. 1 .

Embodiments of the present disclosure can include an oligomerizationcatalyst having transition metal ions selected from the elementsconsisting of chromium, nickel, iron, cobalt, and combinations thereof.The amount of the Group VIII noble metal can range from about 0.001 wt %to about 30 wt %, optionally from about 0.01 wt % to about 10 wt %,optionally from about 1.0 wt % to about 5 wt %, all calculated on anelemental basis of the final catalyst composition. Embodiments of theoligomerization catalyst include nickel present in the form of Ni²⁺cations in an amount ranging from 0.001 wt % to 30 wt %, optionally fromabout 0.01 wt % to about 10 wt %, on an elemental basis of the catalystcomposition within a molecular sieve support having frameworkheteroatoms selected from the group consisting of Al³⁺, Zn²⁺, andcombinations thereof, wherein the Ni²⁺ cations are coordinated with twoframework Al³⁺ heteroatom centers in a paired configuration or with oneframework Zn²⁺ heteroatom center, or combinations thereof. In anembodiment the catalyst is treated for H⁺ sites with the addition ofLi⁺.

Embodiments of the present disclosure can include an oligomerizationcatalyst comprising a crystalline, microporous zeolite with MFI topologyand heteroatoms of aluminum, gallium, or iron within the MFI structurehaving a first bulk heteroatom content and first crystal size. Trivalentboron atoms are added within the support matrixes that have beensubstituted for tetravalent silicon atoms. The selectivity towardoligomers is enhanced compared to catalysts without boron with similarbulk heteroatom content and crystal size. In an embodiment the boron ispresent in an amount ranging from 0.001 wt % to about 10 wt %,optionally from about 0.01 wt % to 5 wt % on an elemental basis of thecatalyst composition. In an embodiment the aluminum, gallium or iron arepresent in an amount from 0.001 wt % to about 10 wt %, optionally fromabout 0.01 to 5 wt % on an elemental basis of the catalyst composition.In an embodiment the substitution of trivalent boron atoms fortetravalent silicon atoms have low activity Brönsted acid sites forcatalysis. In an embodiment the Brönsted acid sites that compensateframework boron atoms are unreactive compared to those that compensateframework aluminum or gallium or iron atoms. In an embodiment theBrønsted acid sites that compensate framework boron atoms aresignificantly weaker acid sites compared to those that compensateframework aluminum or gallium or iron atoms.

The various elements that make up the components for the oligomerizationcatalyst can be derived from any suitable source, such as in theirelemental form, or in compounds or coordination complexes of an organicor inorganic nature, such as carbonates, oxides, hydroxides, nitrates,acetates, chlorides, phosphates, sulfides and sulfonates. The elementsand/or compounds can be prepared by any suitable method known in the artfor the preparation of such materials.

The present disclosure is not limited by the method of oligomerizationcatalyst preparation, and all suitable methods should be considered tofall within the scope herein.

In an embodiment the catalyst of the present disclosure can undergoin-situ regeneration, which can lower operating costs by decreasing theamount of time the reactor must be offline. The regeneration can be doneby hydrogen and water vapor stripping at the reaction temperature. In anembodiment the catalyst of the present disclosure can undergo ex-situregeneration.

An embodiment of the present disclosure is a process for theoligomerization of alkenes to higher molecular weight hydrocarbons. Theprocess includes the steps of introducing an alkene feedstock into areaction chamber, passing the feedstock over an oligomerization catalystat reaction conditions effective to produce oligomerization reactionsconverting a portion of the alkenes to higher molecular weighthydrocarbons. In an embodiment the higher molecular weight hydrocarbonsare substantially free of aromatic hydrocarbons.

The alkene feedstock can be alkenes containing less than 10 carbonatoms. The feedstock can consist primarily of C₂-C₅ alkenes. The alkenefeedstock can be obtained from any suitable source. In an illustrativeembodiment the alkene feed can contain primarily ethene. In anillustrative embodiment the alkene feed can contain primarily propene.In an illustrative embodiment the alkene feed can contain primarilybutene. In an illustrative embodiment the alkene feed can containprimarily ethene and propene. In an illustrative embodiment the alkenefeed can contain primarily propene and butene. In an illustrativeembodiment the alkene feed can contain primarily butene and pentene. Inan illustrative embodiment the alkene feed can contain primarily C₃-C₅alkenes. In an illustrative embodiment the alkene feed can containprimarily C₄-C₅ alkenes.

The reaction chamber can house any suitable catalyst system, such as afixed catalyst bed, a moving bed or a fluidized bed. Single or multiplecatalyst beds can be used, and the reactor can be a swing reactor. Inembodiments the reaction can take place at a temperature of from 50° C.to 350° C. The pressure can be in the range of from 0.5 psig to 600psig. The weight hourly space velocity can be from 0.5 to 600 hr⁻¹.

The reaction products can be processed further if needed, such as with asaturation step to more fully saturate the produced hydrocarbon liquidsand reduce the olefin content. The final reaction products can beseparated by cooling or other standard recovery or separationtechniques.

Process

The NTL process includes converting NGLs to their corresponding olefinsthrough catalytic dehydrogenation. The olefins are then coupled togetherto form high molecular weight mono-olefin hydrocarbons through catalyticoligomerization. The higher molecular weight hydrocarbons can berecovered through a separation system as a liquid product. Thedehydrogenation and oligomerization reactions can be performed inseparate reactors, such as a dehydrogenation reactor and oligomerizationreactor placed in series, or alternatively both the dehydrogenation andoligomerization reactions can be performed within a single combinedreactor. If in a combined reactor the dehydrogenation reaction canproceed the oligomerization reaction by locating a dehydrogenationcatalyst bed followed by an oligomerization catalyst bed. Alternativelythe dehydrogenation catalyst and oligomerization catalyst can both belocated within a mixed bed, wherein the dehydrogenation reaction andoligomerization reaction can each occur within the mixed bed. Followingseparation of the product liquid, the unconverted alkane gas stream isrecycled to the process. The liquid product can undergo furthertreatment as needed, such as a saturation procedure to reduce the olefincontent and make a more saturated liquid hydrocarbon stream suitable forfuel blending. FIG. 2 illustrates one embodiment of this transformationprocess, which is described below.

Referring now to FIG. 2 , the feed stream 10 can be a hydrocarboncontaining gas, which can be a mixture of methane (C₁), C₂, C₃, C₄, andC₅ which can include trace amounts of H₂S, CO₂, N₂, and H₂O. Dependingon the concentration of methane in the feed gas stream 10, methane mayneed to be removed to a low concentration in stream 20 through aseparation system 100, consisting of any suitable separation scheme suchas a typical cryogenic NGL extraction process, distillation and/ormembrane technology. The separated methane stream 90 can then be used asfuel gas or sold as a product. After methane removal, the methane-poorfeed 20 is combined with a recycle stream, C₂₊, 82 which ispredominantly unconverted ethane, propane, ethylene, butylene, andresidual hydrogen. The combined stream 24 is then typically heated to anelevated temperature (650-800° C.) through a heat exchange unit 110 andis passed to a bed of dehydrogenation catalyst within a dehydrogenationreactor 120.

In the dehydrogenation reactor 120, the reactor typically operates atelevated temperature (500-800° C.) and medium pressure (2-25 bar). Theconversions of ethane, propane, butane, and pentane to their olefinscounterparts are endothermic and limited by thermodynamic equilibrium.Although not shown in FIG. 2 , heat is generally provided to thedehydrogenation reactor to increase olefin formation. The supportedcatalyst is designed so that formation of undesirable side products,such as methane and acetylene, is minimized. Bimetallic catalysts can beused as a dehydrogenation catalyst comprised of a Group VIII metal, suchas Pt, Pd, Ni, Co, Fe, Rh, Ir, Ru and a non-catalytic metal such as Zn,Ga, Sn, In, Mn, V, Cr, or Ti. Non-limiting combinations of catalytic andnon-catalytic bimetallic catalysts are PtZn, PdZn, Ptln, Pdln, PtSn, Pt,Mn, PtCr, PtV, PtTi, PdMn, PtCu, and NiIn, which can be employed toachieve the desired reaction. Additionally, the catalysts may includenon-metallic, oxide catalysts of Cr, Co, Zn, Ga, or Ni stabilized byphosphate. The support can include a refractory oxide such as silica,alumina, silica-alumina, titania, magnesium oxide, zeolites andaluminophosphates.

The stream 32 exiting the dehydrogenation reactor 120 contains a mixtureof olefins, unconverted alkanes, and hydrogen. Stream 32 is then cooledto a lower temperature (200-600° C.) through a heat exchange process inheat exchanger 130. Depending on the hydrogen tolerance of thedownstream processes, some portion of the hydrogen must be first removedfrom the cooled stream 40.

Through a hydrogen separation system 140, such as membrane, stripping,and distillation, sufficient hydrogen 42 can be removed from the systemto have only a residual concentration of hydrogen in stream 44, whichmay be needed to ensure catalyst stability and mass balance in theoligomerization reactor 160. The hydrogen separation system 140 shouldpossess high selectivity to hydrogen such that the losses of othercomponents are minimized. After removal of the hydrogen in stream 42,the temperature of stream 44 may be adjusted in heat exchanger 150 andstream 50 is then fed to another reactor 160 with a bed ofoligomerization catalyst.

The oligomerization reactor 160 can operate at a relatively lowtemperature (200-600° C.) and preferably medium pressure (2-50 bar). Inthe oligomerization catalyst bed, the conversion of olefins may not goto completion. The stream 52 exiting this reactor may contain a mixtureof C₁-C₉ with a majority of the higher molecular weight molecules beinggreater than C₆ hydrocarbons, and preferably C₁₀-C₁₈, or highermolecular weight. Catalysts such as H-ZSM₅, and especiallyBoron-modified H-(Al)ZSM-5 can be employed in order to achieve thedesirable reaction.

Stream 52 is then cooled to a lower temperature (2°-10° C.) in heatexchanger 170 to condense the high molecular weight hydrocarbons. Atthis condition, low molecular weight hydrocarbons may also be dissolvedin the condensed liquid. The two-phase stream 60 is then separated in atwo-phase separator 180. The liquid 62 from separator 180 can then bebrought to ambient pressure and ambient temperature through a pressuredrop across a Joule-Thompson value or a dense fluid expander 190 andheat exchange unit 195. This results in a two-phase stream 70 which canthen be fed to another two-phase separator 200. The liquid 72 fromseparator 200 is the product which consists of the high molecular weighthydrocarbon molecules.

The vapor stream 74 from separator 200 contains the unconvertedmolecules such as ethane, propane, butane, ethylene, propylene, andbutylene. Vapor stream 74 from separator 200 is boosted in pressureacross booster 205 and combined with vapor stream 65 from separator 180.Pressure of the vapor from the second flash tank may not need to beincreased through booster 205, instead pressure of stream 64 may bereduced and matched to the pressure of stream 74. Finally, the combinedstream 76, after an optional purge stream 78, is increased in pressureacross recycle compressor 210 to provide a recycle stream 82 which ismixed with the methane-poor stream 20, which results in combined stream24 that subsequently enters the dehydrogenation reactor 120.

Techno-economic analysis of this process on a basis of 96 MMSCFD ofshale gas containing ˜30% mole of NGL from the Bakken field estimated atotal capital cost of 251 million USD with annual operating cost of 19.8million USD producing 15,000 bbl/day of liquid hydrocarbons. AFischer-Tropsch process of the same size is estimated to have a totalcapital cost ranging from 300-525 million USD with an annual operatingcost between 11-28 million USD. This analysis indicates the NTL processis economically lucrative compared to existing technologies.

FIG. 3 shows a variation of the process described in FIG. 2 . Considerthe NTL process shown in FIG. 2 , when methane recovery is not very highin separator 100, the high C₁ content in stream 20 necessitates verylarge recycle flow in stream 80 and 82 leading to high capital andoperating expenses. Therefore, the process in FIG. 3 is designed tohandle poor separation on separator 100. To accommodate for thisvariation, separator 250 is utilized to maximize the recovery of lightalkanes and alkenes from the recycle stream. Separators 250, 180, and200 along with heat exchangers 195, 240 and pressure changers 190, 205,230 constitute a multi-stage separator configuration. Vapor stream 75,instead of being directly recycled after an optional purge, iscompressed in pressure booster 230 followed by cooling to lowtemperature (2°-10° C.) in heat exchanger 240 which results in atwo-phase stream 95. Stream 95 undergoes two phase separation inseparator 250. Vapor stream 96 pressure is increased through compressor211. To avoid methane accumulation in the recycle stream, vapor stream98 exiting compressor 211 can be recycled to the feed stream 10 ordirectly into separator 100 depending on the unit operations that isemployed for separator 100. Liquid stream 97 containing a substantialconcentration (˜30 mole %) of unconverted C₃ and C₄ olefins with minorconcentration of C₂ olefin can be heated, such as to oligomerizationreactor 160 operating temperature, in heat exchange unit 260. Vaporstream 93 exiting heat exchange unit 260 is then combined with stream 50which results in stream 99 that enters the oligomerization reactor 160.The recycle ratio of this process was found to be 46% less than that ofthe process in FIG. 2 in simulations of the two processes. This NTLprocess was found to reduce capital and operating costs by 8% and 26%,respectively, compared to the process described in FIG. 2 .

An embodiment of the present disclosure is a novel process that isuseful when hydrogen tolerant oligomerization catalyst is available.Hydrogen tolerant oligomerization catalyst is defined as a catalyst thatcan perform alkene coupling on alkenes, such as a mixture of C₂, C₃, C₄,C₅ alkenes, with a substantial concentration of hydrogen (>5 molepercent). FIG. 4 illustrates a variation of the process described inFIG. 2 wherein a hydrogen tolerant oligomerization catalyst is employed.The partial pressure of hydrogen in stream 103 exiting thedehydrogenation reactor 120 is lower than the partial pressure of thehydrogen in vapor stream 104 as light alkenes have been both convertedand removed as high molecular weight hydrocarbons by oligomerizationreactor 160 and separator 180. This higher partial pressure of hydrogenin vapor stream 104 can be exploited by performing hydrogen separationwith separator 270 to remove hydrogen via stream 101 and produce ahydrogen lean stream 102. Ultimately, this process has a reduced capitalexpense for separator 270 compared to separator 140 in FIG. 2 .

To minimize heat input into the processes, heat and mass integration canbe performed on the described process. An example of integrated processdesign for the process in FIG. 4 is shown in FIG. 5 . In this example,heat exchangers 110, 130, 170 within the recycle loops are coupled suchthat the remaining heat duty is for the catalytic dehydrogenationreactor 120 that operates at elevated temperature, such as from500°-800° C. Combined stream 24 enters a process-to-process heatexchanger 280 and is heated using stream 108 and stream 106 which exitoligomerization reactor 160 and dehydrogenation reactor 120,respectively. The resulting heated stream 105 enters dehydrogenationreactor 120 and exhaust stream 106 enters heat exchanger 280. Heat fromstream 106 is transferred to stream 24 until stream 106 temperaturereaches the operating temperature of oligomerization reactor 160.Exhaust stream 108 from oligomerization reactor 160 enters heatexchanger 280 and transfers its heat before exiting as stream 109.Finally, stream 109 is cooled to low temperature (˜2°-10° C.) and exitsa two-phase stream 110. There are other integrated processconfigurations that have not been illustrated here. This integratedprocess flowsheet reduces the number of heat exchangers and alsoutilities costs which ultimately lead to reduction in the total capitalcost and annual operating costs.

Further examples of integrated process design can include integrationwith an existing NGL extraction plant, such as by utilizing waste heatfrom the NGL plant to provide a portion of the heat load required forthe dehydrogenation reaction. In a non-limiting example therecompression of residue gas from an NGL extraction plant may beprovided by one or more gas fired turbine driven compressor. The turbineexhaust temperature can be more than 600° C. and can be used as a heatsource within the NTL process.

Hydrogen is a by-product of a dehydrogenation reactor and it can beutilized in several ways depending on the deployment of the NTL process.Where the NTL process is integrated into a large refinery complex, thenthe produced hydrogen can be delivered into processes that require themeither as feedstocks or as fuels. Produced hydrogen can be utilized in asaturation step to more fully saturate the produced hydrocarbon liquidsand reduce the olefin content. In remote locations, hydrogen must beused or transformed on-site as hydrogen pipeline infrastructure is notextensive. Although it is intuitive to utilize hydrogen as fuel gas tosupply heat for dehydrogenation reactor, hydrogen can also be used tofix any available on-site CO₂ into valuable chemicals such as methanoland dimethyl ether which are also liquid at ambient temperature andpressure. This proposed NTL process with hydrogen fixation on CO₂ isdescribed in FIGS. 6 and 7 . In FIGS. 6 and 7 , the hydrogen fixationprocess is integrated with NTL processes described in FIGS. 2 and 3 ,respectively. In FIG. 2 , raw shale gas stream 109 coming out of thewellhead usually contains varying concentrations of CO₂, N₂, H₂S, andH₂O, which may need to be removed before it can enter the natural gaspipeline or undergo further processing. The removal process 290 isgenerally called gas treatment which may entail some or all thefollowing processes: acid gas removal, dehydration, membrane separation,and cryogenic distillation. From gas treatment 290, a CO₂ stream 111 isrecovered and hydrocarbon gas stream 10 is sent for further processingas described in FIG. 2 . A portion of the methane rich stream 92 can beused to supply heat for other unit operations such as a dehydrogenationreactor by combusting the split stream 113 in a furnace 320 with an airstream 114. The exhaust gas stream 115 from furnace 320 enters separator330 to generate a CO₂ rich stream 117. Streams 111 and 117 can then becombined and enter a hydrogen fixation process 340. The hydrogen stream42 that is separated by hydrogen separator 140 can be sent to a hydrogenfixation process 340. Direct hydrogenation of CO₂ to methanol can occurat medium temperature (˜235°-250° C.) and medium pressure (˜5-50 bar)using Cu/Zn catalysts. The resulting outlet stream 119 contains methanoland other products such as higher boiling alcohols and dimethyl ether.This process allows co-production of valuable chemicals and liquidhydrocarbons while reducing the CO₂ emission from the process.

An embodiment of the present disclosure is a process to catalyticallytransform natural gas liquid (NGL) into higher molecular weighthydrocarbons includes providing an NGL stream, catalyticallydehydrogenating at least a portion of the NGL stream components to theircorresponding alkene derivatives, catalytically oligomerizing at least aportion of the alkenes to higher molecular weight hydrocarbons andrecovering the higher molecular weight hydrocarbons. The NGL stream canbe extracted from a gas stream such as a gas stream coming from shaleformations. The higher molecular weight hydrocarbons can be hydrocarbonsthat are liquid at ambient temperature and ambient pressure.

In an embodiment, the process can include a multi-stage separationconfiguration to recover a portion of the light alkenes from a recycleloop and use the light alkenes as feed to the catalytic oligomerization.A hydrogen tolerant oligomerization catalyst can be used and hydrogenseparation can be employed from a recycle stream. The process caninclude process-to-process heat exchangers employed to provide heatrecovery within a recycle loop. The process can optionally include ahydrogen fixation onto CO₂ process employed to generate liquid alcoholproducts, such as methanol, dimethyl ether, and higher boiling alcohols.

The text above describes one or more specific embodiments of a broaderdisclosure. The disclosure also can be carried out in a variety ofalternate embodiments and thus is not limited to those described here.The foregoing description of an embodiment of the disclosure has beenpresented for the purposes of illustration and description. It is notintended to be exhaustive or to limit the disclosure to the precise formdisclosed. Many modifications and variations are possible in light ofthe above teaching. It is intended that the scope of the disclosure belimited not by this detailed description, but rather by the claimsappended hereto.

What is claimed is:
 1. A non-oxidative process for catalyticallytransforming a natural gas liquid (NGL) into one or more highermolecular weight hydrocarbons, comprising: providing a hydrocarboncontaining gas stream comprising C₂ and C₃ alkanes; catalyticallydehydrogenating at least a portion of the hydrocarbon containing gasstream to corresponding alkene derivatives using at least one bimetalliccatalyst, providing an effluent comprising of unreacted C₂-C₃ alkanes,C₂-C₃ alkenes and hydrogen; catalytically oligomerizing at least aportion of the alkenes in the effluent to provide one or more highermolecular weight hydrocarbons; and recovering the higher molecularweight hydrocarbons.
 2. The process according to claim 1, wherein thehydrocarbon containing gas stream is derived from a shale formation. 3.The process according to claim 1, further comprising recycling unreactedalkanes and alkenes after the catalytic oligomerization reaction.
 4. Theprocess according to claim 1, wherein the at least one bimetalliccatalyst comprises a first metal chosen from the group consisting of Pt,Pd, Co, Fe, Rh, Ir, and Ru, and a second metal chosen from the groupconsisting of Zn, Mn, V, and Cr.
 5. The process according to claim 1,wherein the at least one bimetallic catalyst is chosen from the groupconsisting of PtZn, PdZn, PtIn, Pdln, PtMn, PtCr, PtV, and PdMn.
 6. Theprocess according to claim 1, wherein the at least one bimetalliccatalyst comprises a support of refractory oxide chosen from the groupconsisting of silica, alumina, silica-alumina, titania, magnesium oxide,zeolites, aluminophosphates and combinations thereof.
 7. The processaccording to claim 1, wherein the at least one bimetallic catalyst is azero valent metallic alloy that comprises a combination of a first metalchosen from the group consisting of platinum and palladium, and a secondmetal chosen from the group consisting of manganese, vanadium, chromiumand combinations thereof, and a support.
 8. The process according toclaim 1, further comprising removing the hydrogen after catalyticdehydrogenation; and utilizing the removed hydrogen to increasesaturation and reduce the alkene content of the higher molecular weighthydrocarbon stream.
 9. The process according to claim 1, wherein the atleast one bimetallic catalyst comprises a support of refractory oxidechosen from the group consisting of silica, alumina, and zeolites. 10.The process according to claim 1, wherein the hydrocarbon containing gasstream comprising C₂ and C₃ alkanes contains no isobutane.
 11. Anon-oxidative process for catalytically transforming a natural gasliquid (NGL) into one or more higher molecular weight hydrocarbons,comprising: providing a hydrocarbon containing gas stream consisting ofC₂ and C₃ alkanes; catalytically dehydrogenating at least a portion ofthe hydrocarbon containing gas stream to corresponding alkenederivatives and hydrogen using at least one bimetallic catalyst,providing an effluent consisting of unreacted C₂-C₃ alkanes, C₂-C₃alkenes and hydrogen; catalytically oligomerizing at least a portion ofthe C₂-C₃ alkenes in the presence of at least one oligomerizationcatalyst to provide one or more higher molecular weight hydrocarbons;and recovering the higher molecular weight hydrocarbons, wherein the atleast one oligomerization catalyst comprises H-ZSM-5, B—Al—H—ZSM-5, or acombination thereof.
 12. The process according to claim 11, wherein thehydrocarbon containing gas stream is derived from a shale formation. 13.A process for catalytically transforming a natural gas liquid (NGL) intoone or more higher molecular weight hydrocarbons, comprising: providinga hydrocarbon containing gas stream that is derived from a shaleformation, the stream comprising one or more light alkanes chosen fromthe group consisting of C₂, C₃, C₄, C₅, and combinations thereof;catalytically dehydrogenating at least a portion of the light alkanes tocorresponding alkene derivatives; catalytically oligomerizing at least aportion of the alkenes in the presence of at least one oligomerizationcatalyst to provide one or more higher molecular weight hydrocarbons;and recovering the higher molecular weight hydrocarbons, wherein the atleast one oligomerization catalyst comprises: a crystalline microporouszeolite structure having MFI topology; heteroatoms within the MFIstructure selected from the group consisting of aluminum, gallium, ironand combinations thereof; and trivalent boron atoms within the MFIstructure that have been substituted for tetravalent silicon atoms. 14.The process according to claim 13, wherein the boron is present in anamount ranging from 0.001 wt % to 10 wt % on an elemental basis of theoligomerization catalyst.
 15. The process according to claim 13, whereinthe aluminum, gallium or iron are present in an amount from 0.001 wt %to 30 wt % on an elemental basis of the oligomerization catalyst. 16.The process according to claim 13, wherein the substitution of trivalentboron atoms for tetravalent silicon atoms do not behave as Brönsted acidsites for catalysis.
 17. The process according to claim 13, wherein aBrönsted acid site that compensate framework boron atoms have lowactivity compared to those that compensate framework aluminum or galliumor iron atoms.
 18. The process according to claim 13, wherein a Brönstedacid site that compensate framework boron atoms are significantly weakeracid sites compared to those that compensate framework aluminum orgallium or iron atoms.